Method and system for generating hydrogen-enriched fuel gas for emissions reduction and carbon dioxide for sequestration

ABSTRACT

A method of generating hydrogen-enriched fuel gas and carbon dioxide comprising converting hydrocarbon molecules from a gaseous hydrocarbon feed stream into hydrogen and carbon dioxide, separating the hydrogen and carbon dioxide, blending the hydrogen back into the gaseous hydrocarbon feed stream to generate a hydrogen-enriched fuel gas, and utilizing the carbon dioxide for storage or sequestration. A system for generating hydrogen-enriched fuel gas and carbon dioxide comprising an inlet handling system, a syngas and water-gas shift system, a water-gas compression system, a carbon dioxide recovery system, a dehydration system, and a carbon dioxide compression system.

CROSS-REFERENCE TO RELATED APPLICATION

This application is a continuation of U.S. patent application Ser. No.12/101,087, filed on 10 Apr. 2008.

BACKGROUND OF THE INVENTION

1. Field of the Invention

The present invention relates generally to the field of oil and gas, andmore specifically, to a method and system for generatinghydrogen-enriched fuel gas for emissions reduction and carbon dioxidefor sequestration.

2. Description of the Related Art

Climate change resulting in global warming from increased greenhouse gas(GHG) concentration in the atmosphere has been a problem for many years,and the recent accelerated increase in GHG concentration in theatmosphere and the associated climatic impacts identified around thework has been alarming. The anthropogenic (man-made) causes of globalwarming include carbon dioxide emissions into the atmosphere fromhydrocarbon fuel combustion, including, but not limited to, electricalpower generation facilities, vehicles, railroad locomotives, ships,airplanes, petrochemical, manufacturing, industrial and commercialindustries; methane derived from agricultural sources (e.g., ricepaddies, bovine flatulence, etc.), biomass (human waste, animal waste,and agricultural waste), and hydrocarbon fuel production; increasedwater vapor in the atmosphere from hydrocarbon fuel combustion; nitrousoxide from hydrocarbon fuel combustion; and deforestation, whichreleases hundreds of millions of tons of carbon dioxide into theatmosphere each year. [1] According to the United NationsIntergovernmental Panel on Climate Change (IPCC), it is projected thatglobal warming will cause dry areas to get drier, drought-affected areasto become larger, an increase in heavy precipitation events, and adecrease in water supply stored in glaciers and snow pack. In addition,coastal areas will be exposed to coastal erosion and sea-level rise, andmany millions more people will be flooded every year due to sea-levelrise by the 2080s. [2]

Hydrocarbon fuel combustion emissions are a major contributor to GHGconcentrations that impact climate change and result in global warming.It has been documented that hydrogen-enriched fuel decreases emissions.[3, 4] The problem that has not been solved, however, is how to produceand develop a hydrogen-based infrastructure. Most of the debatesurrounding this issue involves developing a pure hydrogen product thateliminates carbon dioxide emissions from combustion of the fuel.Developing a blended hydrogen and pipeline quality natural gas productthat utilizes the existing pipeline transportation and distributionsystem would reduce emissions from industrial, commercial, residentialand mobile natural gas consumers. This problem presents challenges thatheretofore have not been solved by the prior art. The present inventionoffers a solution to this problem by providing a method and system:converting the hydrocarbon molecules to hydrogen and carbon dioxide;separating the hydrogen and carbon dioxide; storing, sequestering orutilizing the carbon dioxide so that it is not emitted into theatmosphere; and blending the hydrogen back into the natural gas and/orusing the hydrogen as fuel. In a preferred embodiment, a carbon dioxiderecovery solvent is used to separate the carbon dioxide from thehydrogen, which has the advantage of allowing the separation step todeliver the hydrogen and carbon dioxide at high pressure, therebyreducing the cost and energy usage of compressing and transportingcarbon dioxide downstream systems and the hydrogen back to the naturalgas pipeline. The carbon dioxide recovery solvent also significantlylowers circulation rates when compared to conventional physical andchemical solvents, thereby reducing the energy required and size ofequipment needed to implement the process.

The term “carbon sequestration” generally refers to the long-termstorage of carbon in a multitude of means, including, but not limitedto, terrestrial, underground, or ocean environments to reduce thebuildup of carbon dioxide in the atmosphere. [5] One method ofcontaining carbon dioxide—called “geologic sequestration”—is to injectit into geologic formations (e.g., coal beds, petroleum formations,saline aquifers, basalt formations, etc.). Enhanced oil recovery (EOR),a form of geologic sequestration, is the process by which carbon dioxideand sometimes water are injected into oil reservoirs, thereby flushingout reserves of oil that would otherwise remain unrecovered. Thisprocess can extend the life of an oil reservoir by years, and in somecases, produce many millions of barrels of extra oil without causingsubstantial additional impacts to the surface and eliminating ordelaying the expansion of oil exploration into sensitive areas. [6] In apreferred embodiment of the present invention, the carbon dioxide thatis recovered as part of the present invention is used in EOR operations.

Current technologies for capturing carbon from fossil fuels generallyfall into two categories: pre-combustion and post-combustion. To date,pre-combustion technologies have been limited to removal of carbondioxide from coals, natural gas or syngas immediately prior tocombustion. Though many plants have been proposed to deliver thecaptured carbon dioxide for eventual sequestration, no commercial plantshave been constructed to date. Prior art post-combustion technologiesare limited to removing carbon dioxide from low-pressure (nearatmospheric) flue gases using chemical amine-based solvents or chilledammonium bicarbonate solvent. These applications typically require highsolvent circulation rates, high thermal energy requirements, and oftenhave problems caused by oxygen contamination and high flue gastemperatures. In addition, the delivery pressure of the carbon dioxidefrom said prior art systems is limited to 5-10 psig (pound-force persquare inch gauge), which requires high energy consumption due tocompressing the carbon dioxide for sequestration in geological storagereservoirs or for use for carbon dioxide EOR.

Furthermore, much of the prior art deals with carbon capture andsequestration at point sources (e.g., individual stationary facilities);however, most point sources generate relatively small amounts of carbondioxide that are not of sufficient quantities to support EOR and othertypes of geologic sequestration. To overcome this limitation, carbondioxide compression and pipeline networks need to be constructed toaggregate carbon dioxide volumes from point sources, which is onlypractical for large industrial point sources. The present inventionavoids this problem by generating carbon dioxide not at a point sourcebut at a strategically located hydrogen generation plant on the mainpipeline (well upstream of the point sources), preferably located in anarea that would both geologically support either carbon dioxidestorage/sequestration or carbon dioxide EOR injection. In this manner,sufficient quantities of carbon dioxide can be produced to make carboncapture and storage/sequestration economically feasible.

BRIEF SUMMARY OF THE INVENTION

The present invention is a method of generating hydrogen-enriched fuelgas and carbon dioxide comprising: converting hydrocarbon molecules froma gaseous hydrocarbon feed stream into hydrogen and carbon dioxide;separating the hydrogen and carbon dioxide; blending the hydrogen backinto the gaseous hydrocarbon feed stream to generate a hydrogen-enrichedfuel gas; and utilizing the carbon dioxide for storage or sequestration.In a preferred embodiment, each standard cubic foot of the gaseoushydrocarbon feed stream produces between two and four standard cubicfeet of a hydrogen product stream and between 0.7 and 0.9 standard cubicfeet of a carbon dioxide product stream. Preferably, thehydrogen-enriched fuel gas has a hydrogen concentration ranging fromfive to 30 mole percent.

In a preferred embodiment, the hydrogen-enriched fuel gas produces lesscarbon dioxide per energy unit output when combusted thannon-hydrogen-enriched natural gas. Preferably, approximately one toeleven percent of the gaseous hydrocarbon feed stream is processed.

In a preferred embodiment, the total volume of the gaseous hydrocarbonfeed stream ranges from 100 million standard cubic feet per day to 4500million standard cubic feet per day, and between 10 million standardcubic feet per day and 500 million standard cubic feet per day of thegaseous hydrocarbon feed stream is processed. Preferably, thehydrogen-enriched fuel gas is transported and distributed using theexisting natural gas pipeline system. In a preferred embodiment, thesequestration is enhanced oil recovery.

In a preferred embodiment, the carbon dioxide is separated from thehydrogen using a carbon dioxide recovery solvent, and the carbon dioxiderecovery solvent is one or more hydrocarbon liquid(s) selected from thegroup consisting of butanes, pentanes, hexanes, heptanes, octanes,aromatics, and isomers of butanes, pentanes, hexanes, heptanes, octanesand aromatics. Preferably, the carbon dioxide recovery solvent isnormal-butane or a mixture of normal-butane and iso-butane. In apreferred embodiment, the carbon dioxide recovery solvent allows thecarbon dioxide to be separated from the hydrogen at a pressure ofbetween 200 and 500 psig.

In a preferred embodiment, total carbon dioxide compression requirementsfor the storage or sequestration are reduced by 50 to 75 percent ascompared to chemical or physical solvent-based carbon dioxide recoveryprocesses that do not utilize the carbon dioxide recovery solvent of thepresent invention. Preferably, the conversion, separation and blendingsteps occur on a natural gas transportation and distribution pipelineand not at a point of combustion. In a preferred embodiment, theconversion, separation and blending steps are all performed prior tocombustion of the hydrogen-enriched fuel gas.

In a preferred embodiment, the conversion, separation and blending stepsdo not require any changes to the existing natural gas pipelinetransportation and distribution system other than providing mobilepoints of consumption with an ability to consume the hydrogen-enrichedfuel gas and increasing the number of compressed natural gas fuelingfacilities to supply the mobile points of consumption with thehydrogen-enriched fuel gas. Preferably, the gaseous hydrocarbon feedstream is pipeline quality natural gas. In a preferred embodiment, thepresent invention further comprises utilizing a portion of the separatedhydrogen as a separate fuel product.

In an alternate embodiment, the present invention is a method ofgenerating hydrogen and carbon dioxide comprising: convertinghydrocarbon molecules from a gaseous hydrocarbon feed stream intohydrogen and carbon dioxide; separating the hydrogen and carbon dioxide;utilizing the hydrogen as a separate product; and utilizing the carbondioxide for storage or sequestration; wherein the carbon dioxide isseparated from the hydrogen using a carbon dioxide recovery solvent; andwherein the carbon dioxide recovery solvent is one or more hydrocarbonliquid(s) selected from the group consisting of butanes, pentanes,hexanes, heptanes, octanes, aromatics, and isomers of butanes, pentanes,hexanes, heptanes, octanes and aromatics. Preferably, the carbon dioxiderecovery solvent is normal-butane or a mixture of normal-butane andiso-butane. In a preferred embodiment, the carbon dioxide recoverysolvent allows the carbon dioxide to be separated from the hydrogen at apressure of between 200 and 500 psig.

The present invention is also a system for generating hydrogen-enrichedfuel gas and carbon dioxide comprising: an inlet handling system; asyngas and water-gas shift system; a water-gas compression system; acarbon dioxide recovery system; a dehydration system; and a carbondioxide compression system; wherein the inlet handling system prepares agaseous hydrocarbon feed stream having a pressure and a temperature tobe fed to the syngas and water-gas shift system by removing liquids andsolids, reducing the pressure of the gaseous hydrocarbon feed stream,and maintaining the temperature of the gaseous hydrocarbon stream fordownstream processes; wherein the syngas and water-gas shift systemproduces hydrogen-bearing syngas by reforming the gaseous hydrocarbonfeed stream into hydrogen and carbon monoxide and then converts thecarbon monoxide to carbon dioxide, thereby producing a water-gas stream;wherein the carbon dioxide recovery system has an operating pressure;wherein the water-gas compression system compresses the water-gas streamfrom the syngas and water-gas shift system to the operating pressure ofthe carbon dioxide recovery system; wherein the carbon dioxide recoverysystem separates the carbon dioxide from the hydrogen contained in thewater-gas stream produced by the syngas and water-gas shift system;wherein the dehydration system removes water vapor present in thewater-gas stream produced by the syngas and water-gas shift system,thereby recovering additional carbon dioxide; and wherein the carbondioxide compression system compresses the carbon dioxide recovered fromthe water-gas stream by the carbon dioxide recovery system and thedehydration system to a carbon dioxide delivery pressure.

In a preferred embodiment, the carbon dioxide that is compressed by thecarbon dioxide compression system is used in a particular sequestrationprocess that has a required delivery pressure, and the carbon dioxidedelivery pressure is determined by the required delivery pressure of thesequestration process for which the carbon dioxide is used. Preferably,the sequestration process is enhanced oil recovery. In a preferredembodiment, the gaseous hydrocarbon feed stream is pipeline qualitynatural gas.

In a preferred embodiment, the carbon dioxide recovery system separatesthe carbon dioxide from the hydrogen contained in the water-gas streamproduced by the syngas and water-gas shift system using a carbon dioxiderecovery solvent that is one or more hydrocarbon liquid(s) selected fromthe group consisting of butanes, pentanes, hexanes, octanes, aromatics,and isomers of butanes, pentanes, hexanes, heptanes, octanes andaromatics. Preferably, the carbon dioxide recovery solvent isnormal-butane or a mixture of normal-butane and iso-butane. In apreferred embodiment, the carbon dioxide recovery solvent allows thecarbon dioxide to be separated from the hydrogen at a pressure ofbetween 200 and 500 psig.

In an alternate embodiment, the present invention is a system forgenerating hydrogen-enriched fuel gas and carbon dioxide comprising:means for converting hydrocarbon molecules from a gaseous hydrocarbonfeed stream into hydrogen and carbon dioxide; means for separating thehydrogen and carbon dioxide; means for blending the hydrogen back intothe gaseous hydrocarbon feed stream to generate a hydrogen-enriched fuelgas; and means for utilizing the carbon dioxide for storage orsequestration.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a flow diagram illustrating the overall process of the presentinvention, including the six systems that comprise the presentinvention.

FIG. 2 is a flow diagram of the inlet handling system of the presentinvention.

FIG. 3 is a flow diagram of the synthesis gas and water-gas shift systemof the present invention.

FIG. 4 is a flow diagram of the water-gas compression system of thepresent invention.

FIG. 5 is a flow diagram of a first part (gas absorption) of the carbondioxide recovery system of the present invention.

FIG. 6 is a flow diagram of a second part (partial rich solventregeneration and lean solvent cooling) of the carbon dioxide recoverysystem of the present invention.

FIG. 7 is a flow diagram of a third part (solvent regeneration) of thecarbon dioxide recovery system of the present invention.

FIG. 8 is a diagram of the refrigeration system used to aid in theabsorption of carbon dioxide in connection with the carbon dioxiderecovery system of the present invention.

FIG. 9 is a flow diagram of the carbon dioxide compression system of thepresent invention.

REFERENCE NUMBERS

-   1 Inlet filter separator-   2 Feed gas heater-   3 Pressure let-down device-   4 Flue gas stack-   5 Induced draft combustion air blower-   6 Desulfurization heater-   7 Combustion air pre-heater-   8 Pre-reformer pre-heater-   9 Reformer pre-heater-   10 Flue gas cooler-   11 High-temperature shift reactor feed cooler-   12 High-temperature shift reactor-   13 First low-temperature shift reactor feed cooler-   14 Second low-temperature shift reactor feed cooler-   15 Low-temperature shift reactor-   16 Product gas cooler-   17 Desulfurization bed-   18 Pre-reformer-   19 Reformer-   20 Water-gas compression inlet cooler-   21 Water-gas compression first stage suction scrubber-   22 Water-gas compression first stage compressor-   23 Water-gas compression first stage inter-cooler-   24 Wager-gas compression second stage suction scrubber-   25 Water-gas compression second stage compressor-   26 Water-gas compression second stage cooler-   27 Water-gas compression discharge knock-out drum-   28 Gas/gas exchanger-   29 Presaturator separator-   30 Presaturator pump-   31 High-level gas chiller-   32 Low-level gas chiller-   33 Presaturator chiller-   34 Glycol separator-   35 Ethylene glycol exchanger-   36 Primary absorber-   37 Flash gas compressor-   38 Flash gas compressor suction scrubber-   39 Flash absorber-   40 Low-level solvent chiller-   41 High-level solvent chiller-   42 Ethylene glycol heater-   43 Ethylene glycol flash separator-   44 Solvent cross exchanger-   45 Solvent charge pump-   46 Solvent stripper-   47 Solvent stripper reflux pump-   48 Solvent stripper reboiler-   49 Solvent stripper reflux condenser-   50 Solvent stripper reflux accumulator-   51 Refrigerant subcooler-   52 Refrigerant condenser-   53 Refrigerant compressor first stage suction scrubber-   54 First stage refrigerant compressor-   55 Second stage refrigerant compressor-   56 Refrigerant second stage economizer-   57 Refrigerant first stage economizer-   58 Refrigerant accumulator-   59 Carbon dioxide compressor suction scrubber-   60 Carbon dioxide compressor-   61 Carbon dioxide compressor discharge cooler-   62 Third stage refrigerant compressor

As used in the figures, “LP” means low-pressure. “HP” meanshigh-pressure. “BFW” means boiler feed water. “CW” means cooling water.“EG” means ethylene glycol. “ATM” means atmosphere.

DETAILED DESCRIPTION OF INVENTION

A. Overview

The present invention is a method and system by which a portion of agaseous hydrocarbon feed stream is converted into hydrogen and carbondioxide products. In a preferred embodiment, each standard cubic foot(SCF) of gaseous hydrocarbon feed produces between two and four SCF of ahydrogen product stream and between 0.7 and 0.9 SCF of a carbon dioxideproduct stream. The hydrogen product is then blended with the balance ofthe gaseous hydrocarbon stream to produce a hydrogen-enriched fuel gas.In an alternate embodiment, pure hydrogen is also a product. The carbondioxide product stream is in a form suitable for storage orsequestration.

In a preferred embodiment, the hydrogen-enriched fuel gas streamproduced by the present invention has a hydrogen concentration rangingfrom five to 30 mole percent. To achieve these blends, approximately oneto eleven percent of the total gaseous hydrocarbon stream is processed.The present invention is preferably designed to process between 10million standard cubic feet per day (MMSCFD) and 500 MMSCFD of a gaseoushydrocarbon feed stream out of a total gaseous hydrocarbon streamranging between 100 MMSCFD and 4500 MMSCFD, respectively, to obtain thedesired enriched fuel gas compositions.

FIG. 1 provides an overview of the present invention. The presentinvention is a method and system designed to separate a portion of anatural gas or methane-rich stream; to convert the smaller portion ofthese streams into a mixed hydrogen/carbon dioxide stream; to separatethe hydrogen from the carbon dioxide using a carbon dioxide recoverysolvent whose physical properties allow the delivery pressure of thecaptured carbon dioxide to be uniquely high without the use of gascompression; to take advantage of this high carbon dioxide deliverypressure to reduce the total carbon dioxide compression requirements forEOR or geological storage/sequestration by 50 to 75 percent whencompared to prior art chemical or physical solvent-based carbon dioxiderecovery processes; and to reblend the resulting hydrogen-rich streamwith larger portion of the original natural gas or methane-rich stream,making a gas that will produce less carbon dioxide per energy unitoutput when combusted than the original natural gas or methane-richstream would produce.

The present invention is comprised of six separate systems. The inlethandling system prepares the gaseous hydrocarbon stream to be fed to thesyngas and water-gas shift system by removing liquids and solids,reducing the pressure of the gaseous hydrocarbon stream, and maintainingthe temperature of the gaseous hydrocarbon stream for downstreamprocesses. The syngas and water-gas shift system produceshydrogen-bearing syngas by the reforming of a gaseous hydrocarbon feedstream into mainly hydrogen and carbon monoxide and then “shifting” thecarbon monoxide to carbon dioxide (the mixture commonly known as“water-gas”). The water-gas compression system compresses the resultingwater-gas stream from the syngas and water-gas shift system to theoperating pressure of the carbon dioxide recovery system. The carbondioxide recovery system separates the carbon dioxide from thehydrogen-rich synthesis/shifted gas stream. The dehydration systemremoves the water vapor remaining from the production of the water-gasstream (by the syngas and water-gas shift system) to a level that isacceptable for the carbon dioxide recovery system and the productspecifications of the hydrogen-enriched gas stream and the carbondioxide product. The carbon dioxide compression system compresses therecovered carbon dioxide product streams from both the carbon dioxiderecovery system and the dehydration system to the final carbon dioxidedelivery pressure as required for storage/sequestration in geologicalstorage reservoirs or for use for carbon dioxide EOR.

The present invention is described more fully below in connection withthe six systems that comprise the present invention: inlet handling,synthesis gas (syngas) and water-gas shift, water-gas compression,carbon dioxide recovery, dehydration, and carbon dioxide compression.

B. Inlet Handling System

FIG. 2 is a flow diagram of the inlet handling system of the presentinvention. The inlet handling system prepares the gaseous hydrocarbonstream to be fed to the syngas and water-gas shift system by removingliquids and solids, and both reducing the pressure and maintaining thetemperature of said gaseous hydrocarbon stream for downstream processes.The present invention is designed to accept hydrocarbon feeds withcarbon dioxide concentrations of between 0 and 20 CO₂ mole percent; withhydrogen sulfide concentrations of between 0 and 1 grain per 100standard cubic feet (SCF); and with water vapor contents of between 0%and full saturated. In a preferred embodiment, other gas qualities canbe handled by installing traditional gas processing and conditioningprocesses that enable a wider range of gas compositions, and the presentinvention can be modified to address a variety of conditions. The rangeof operating pressures of acceptable feed gas streams is from 0 psig to2200 psig. The range of temperatures of acceptable feed gas streams isfrom ambient conditions to 300° F.

Referring to FIG. 2, a gaseous hydrocarbon stream is split into agreater stream and a lesser stream. The greater stream flows to theoutlet of the present invention as the gaseous hydrocarbon bypassstream, while the lesser stream, the gaseous hydrocarbon feed, flows tothe inlet handling system of the present invention.

The inlet handling system is comprised of an inlet filter separator 1, afeed gas heater 2, and a pressure let-down device 3. The inlet filterseparator 1 removes any potential free liquids (such as water,hydrocarbons and other impurities such as hydrate inhibitors (methanol,ethylene glycol and tri-ethylene glycol), lube oil, corrosioninhibitors, processing chemicals (amine, potassium, carbonates, etc.),etc.) and solid particulates (such as mill scale, rust, welding slag,dirt, sand, dust, etc.) in the gaseous hydrocarbon feed prior to the gasflowing to the feed gas heater 2. Any accumulated liquids are dumped toa high-pressure drain system for disposal or recycling.

The filtered gaseous hydrocarbon feed then flows to the feed gas heater2. The feed gas heater cross-exchanges the feed gas with low-pressuresteam (50 psig) or other heat medium. The steam or other heat mediumflows to the gaseous hydrocarbon feed gas heater 2, where either latentheat (steam that is condensed) or sensible heat (heat medium that iscooled while the feed gas is heated) is used to increase the temperatureof the feed gas. When steam is used as the heat medium, the steamcondensate flows from the feed gas heater 2 to a feed heater condensatepot (not shown), where it is collected and recycled to a low-pressurecondensate system (not shown) for reuse in steam production. The outletgas temperature of the feed gas from the feed gas heater 2 is controlledto 260° F. or other temperature necessary to maintain the heat of thefeed gas downstream of the pressure let-down device. The heated gaseoushydrocarbon feed then flows to the inlet of the pressure let-down device3, where the gas pressure is reduced from approximately 1220 psia(pounds-force per square inch absolute) to 420 psia. The pressurelet-down device may be an expansion valve (also known as aJoules-Thomson of JT valve) and/or expansion turbine (turbo-expander).Turbo-expanders are isentropic machines that may includeexpander/compressors, expander/generators, and expander/brakes. Theenergy from the turbo-expander is utilized to gain higher efficiencywithin the processing plant by recovering the pressure energy as eithermechanical energy or to generate electrical power.

The inlet handling system described above and the synthesis gas andwater-gas shift system described below are both based upon steamreforming. The present invention, however, is not limited to anyparticular type of reforming method. If an alternative reformingtechnology is used, such as autothermal reforming or partial oxidationreforming, the pressure let-down device 3 can be used to supply thelow-temperature cooling energy for use in an air separation unit orother heat integration methods. The air separation unit would be used toobtain a nearly pure oxygen feed stream in the alternate reformerreaction. The pressure let-down device 3 could also be used as a powersource for a separate piece of rotating equipment, such as a large pumpor compressor anywhere within the present invention. Alternately, thepressure let-down device 3 could be eliminated and replaced with anexpansion valve such as a control valve or choke.

C. Synthesis Gas and Water-Gas Shift System

FIG. 3 is a flow diagram of the synthesis gas and water-gas shift systemof the present invention. In a preferred embodiment, the syngas andwater-gas shift system is based on steam reforming, but the presentinvention is not limited to this particular method. Alternatetechnologies, such as autothermal reforming and partial oxidationreforming, could be used in lieu of steam reforming to create syngas(also called synthesis gas, which is a mixture of mainly hydrogen andcarbon monoxide with other compounds including nitrogen, methane, carbondioxide and water vapor). All of these methods—steam reforming,autothermal reforming and partial oxidation reforming—are known to thoseskilled in the art of petrochemical engineering.

The syngas produces hydrogen by the catalytic oxidation of a gaseoushydrocarbon feed stream. With steam reforming, the hydrocarbon gasstream is heated, mixed with steam and passed over the catalytic beds ofa pre-reformer and a reformer. At the high operating temperatures of thepre-reformer and reformer, the hydrocarbons will react with steam in thepresence of an industry standard, conventional metal-based reformingcatalyst to produce hydrogen, carbon monoxide, carbon dioxide and water.The excess energy produced from the furnace of the reformer is recoveredand used to produce high-pressure steam; to heat feed streams to thepre-reformer and reformer; to preheat combustion air; and/or to supplyenergy to other heating services as needed. The water-gas shift reactionconverts carbon monoxide to carbon dioxide.

Referring to FIG. 3, the syngas and water-gas shift system is comprisedof a desulfurization heater 6, two desulfurization beds 17, apre-reformer pre-heater 8, a pre-reformer 18, a reformer pre-heater 9, areformer 19, a flue gas cooler 10, a high-temperature shift reactor(HTSR) feed cooler 11, an HTSR 12, a first low-temperature shift reactor(LTSR) cooler 13, a second LTSR cooler 14, an LTSR 15, a product gascooler 16, a combustion air pre-heater 7, an induced draft combustionair blower 5, and a flue gas stack 4.

Gaseous hydrocarbon feed from the inlet handling system enters thesyngas and water-gas system at 420 psia and 132° F. The feed stream isheated to 350° F. in the desulfurization heater 6. This is the optimumfeed temperature for the desulfurization units 17. The gaseoushydrocarbon feed then enters the desulfurization units 17, where anytrace amounts of sulfur compounds in the feed stream are removed.Desulfurization of the gaseous hydrocarbon feed is critical becausesulfur compounds will poison the metal-based reforming catalyst used inthe pre-reformer 18 and reformer 19. In one embodiment, thedesulfurization reaction is defined as follows:

H₂S_((v))+Z_(n)O_((s))→Z_(n)S_((s))+H₂O_((v))

Other desulfurization technologies may be used, however.

High-pressure steam is then mixed with the desulfurized gaseoushydrocarbon feed and pre-heated in the pre-reformer pre-heater 8 to1,000° F. In the pre-reformer 18, any high molecular weight hydrocarbonsare reduced to methane, and a significant portion of the steam-reformingreaction occurs. The present invention is not limited to methanereforming, and other gaseous hydrocarbon feeds may be utilized in thereformer. The steam-reforming reaction is defined as follows:

C_(n)H_(m(v)) +nH₂O+heat→nCO_((v))+(m/2+n)H₂

The partial conversion (steam-reforming reaction) occurs in thepre-reformer 18 at a lower temperature (approximately 1000° F. ascompared to approximately 1500° F.) than would otherwise be required inthe reformer 19. This lower temperature allows the use of less expensiveequipment for the pre-reformer and allows the primary reformer 19 (themost costly item in the syngas system) to be smaller. The effluent gasfrom the pre-reformer 18 is mixed with additional high-pressure steamand is heated in the reformer pre-heater 9 to 1,115° F. The resultingheated process gas stream then flows to the primary reformer 19. In thereformer 19, the steam-reforming reaction continues to near equilibriumcompletion.

The heated process gas stream flows downward through catalyst-filledtubes in the reformer 19. An external top-fired burner in the reformer19 provides the energy necessary for the steam-reforming reaction tooccur in the catalyst-filled tubes. The top-fired unit providesco-current flow for the process gas within the tubes, as well as for theflue gas contained within the flue gas section of the reformer 19. Thisallows the highest flue gas temperature to be flowing in the location ofthe coolest in-tube process gas, and it also allows the lowest flue gastemperature to occur where the in-tube process gas is the hottest. Thisarrangement provides uniform tube-wall temperatures over the entire tubelength. The present invention is not limited to the steam-reformingmethod described above, however, and other reformer layouts are includedwithin the scope of the present invention.

The reformer burner combusts either gaseous hydrocarbon fuel orhydrogen-enriched fuel with pre-heated combustion air. The facility ispreferably designed such that once hydrogen production comes on-line,the reformer fuel feed will be switched from the gaseous hydrocarbonfeed required to start up the plant to the produced hydrogen-enrichedgaseous stream and/or hydrogen produced by the present invention. Thecombustion air pre-heater 7 heats the combustion air (pulled in from theatmosphere by the induced draft combustion air blower 5 to 280° F. priorto the combustion air being fed to the burner of the reformer 19.

The hot effluent flue gas from the burner of the reformer 19 providesthe heat input required for the following heaters: the desulfurizationheater 6, the combustion air pre-heater 7, the pre-reformer pre-heater8, the reformer pre-heater 9, and the flue gas cooler 10. After leavingthe reformer, the flue gas then enters the induced draft combustion airblower 5. The induced draft combustion air blower 5 provides the motiveforce for the combustion air to flow through the reformer 19. From theinduced draft combustion air blower 5, the flue gas then enters thebottom of the flue gas stack 4, where the flue gas is vented toatmosphere. The present invention is not limited to this particularcombustion air and flue gas arrangement, however, and other process heatintegration layouts are included within the scope of the presentinvention.

The effluent process gas from the reformer 19 enters the HTSR feedcooler 11, where it is cooled from 1588° F. to 700° F. The energyreleased from the reactor effluent is used to produce high-pressuresteam. The steam produced from this cooler is at 450 psig.

The cooled stream is at the optimum feed temperature for thehigh-temperature shift reaction that occurs in the HTSR 12. In the HTSR12, the water-gas shift reaction occurs. This reduces the quantity ofcarbon monoxide present in the reformer effluent stream, therebyincreasing the hydrogen and carbon dioxide yield of the process. Thewater-gas shift reaction is defined as follows:

CO_((v))+H₂O_((v))→CO_(2(v))+H_(2(v))+heat

The effluent from the HTSR 12 is cooled in the first LTSR feed cooler13, where the effluent is cooled from 813° F. to 415° F. The energyreleased from cooling the gas stream is used to generate 450 psig steam(steam pressure may range between 300 psig and 600 psig). The gas streamis further cooled to 385° F. in the second LTSR feed cooler 14 toproduce 50 psig steam (steam pressure may range between 25 and 75 psig).The cooled gas stream is fed to the LTSR 15, where the water-gas shiftreaction continues, thus further increasing the hydrogen yield. Thesyngas effluent from the LTSR 15 is cooled in the product gas cooler 16,where the syngas effluent is cooled from 423° F. to 320° F. The presentinvention is not limited to this particular process for cooling thesyngas, producing water-gas, and cooling the water-gas, however, andother heat integration arrangements are included within the scope of thepresent invention.

The cooled syngas stream then flows to the water-gas compression system.Alternately, the carbon dioxide could be recovered from the syngasbefore the syngas is compressed, if the total equipment capital costsand system operating costs for both the water-gas compression system andthe carbon dioxide recovery system would be lower, or if the water-gascompression system were located remotely. In this situation, the syngasstream would be cooled from 320° F. to 90-120° F. (depending upon thetype of cooling used—cooling water or air cooler) and flow to the carbondioxide recovery system before flowing to the water-gas compressionsystem.

D. Water-Gas Compression System

FIG. 4 is a flow diagram of the water-gas compression system of thepresent invention. The water-gas compression system compresses thewater-gas (mixture of mainly hydrogen and carbon dioxide) stream fromthe syngas and water-gas shift system to the operating pressure of thecarbon dioxide recovery system.

The discharge pressures of the water-gas compression system may rangefrom 500 psig to 2200 psig, depending upon the selected operatingpressure of the carbon dioxide recovery system and/or the requiredhydrogen, hydrogen-enriched fuel gas, or carbon dioxide deliverypressures. The final product stream temperature may range from 90° F. to120° F., depending upon the type of cooling used (e.g., cooling water orair cooler) and sales contract quality specifications. The type ofcompressors (i.e., centrifugal, reciprocating, etc.) used will beselected based upon economic considerations, process requirements (e.g.,gas stream flow rates and compression requirements), equipment sizing,and manufacturer selection and pricing. The energy required by thissystem is a function of the final design and configuration and can beprovided by electric motors, hydrogen combustion, hydrogen-enriched fuelgas, or other hydrocarbon fuels.

Referring to FIG. 4, the water-gas compression system is comprised of aninlet cooler 20, a first stage suction scrubber 21, a first stagecompressor 22, a first stage inter-cooler 23, a second stage suctionscrubber 24, a second stage compressor 25, a second stage cooler 26, anda discharge knock-out drum 27. Whether single-stage or multiple-stagecompression is used is dependent upon the overall compression ratio,which is determined by the operating pressure of the reformer systemthat establishes the suction pressure to the water-gas compressionsystem and the outlet pressure that established the discharge pressureof the water-gas compression system. The present invention is notlimited to two stages of compressions.

The first step in the water-gas compression system is that the syngasproduct stream from the syngas and water-gas shift system is cooled to90° F. in the inlet cooler 20. (This temperature may be up to 120° F. ifan air cooler is used instead of cooling water). In a preferredembodiment, the inlet cooler 20 condenses approximately 842 gpm (gallonsper minute) of water vapor from the inlet gas stream.

Water vapor is one of the components present in the water-gas productstream. The inlet cooling of this stream prior to compression condensesthe bulk of the water vapor in the water-gas to liquid water. Theresulting two-phase (gas/water) stream then flows to the first stagesuction scrubber 21, where the free water is separated from the gasstream. This minimizes mechanical/corrosion issues and the amount ofwater that may need to be removed in the downstream processing. Theinlet cooling also reduces the compression power requirements for thefirst stage compressor of the water-gas compression system.

After inlet cooling and liquid water separation, the water-gas from thefirst stage suction scrubber 21 flows to the first stage compressor 22,where the stream is compressed from 300 psig to 665 psig. (Dependingupon the process configuration, this may or may not be the actualinter-stage pressure.) The discharge temperature is 267° F. Thepressurized stream is then cooled in the first stage inter-cooler 23,where the stream is cooled to 90° F. (This temperature may be up to 120°F. if an air cooler is used instead of cooling water). Additional wateris condensed in this cooler and removed in the second stage suctionscrubber 24.

The gas from the second stage suction scrubber 24 flows to the secondstage compressor 25, where it is compressed from 660 psig to between1280 psig and 2200 prig, which increases the temperature of the gas isincreased to between 272° F. and 350° F. The compressed gas stream isthen cooled to 90° F. (this temperature may be up to 120° F. if an aircooler is used instead of cooling water) in the second stage cooler 26,condensing a small amount of water from the water-gas. This liquid wateris removed in the discharge knock-out drum 27. The gas stream from thedischarge knock-out drum flows to the carbon dioxide recovery system.

The condensed water streams from the first stage suction scrubber 21,the second stage suction scrubber 24, and the discharge knock-out drum27 are collected and sent to a process water system (not shown). Thewater streams separated from the water-gas stream by these scrubbers arecombined and preferably treated for use as make-up boiler feed water.

E. Carbon Dioxide Recovery System

FIGS. 5, 6 and 7 are flow diagrams of the first, second and third partsof the carbon dioxide recovery system of the present invention. Thecarbon dioxide recovery system is designed to separate the carbondioxide from the hydrogen-rich water-gas stream. This is accomplished byabsorbing the carbon dioxide in a carbon dioxide recovery solvent,resulting in a hydrogen-rich gas stream at high pressure. This solventis a mixture of hydrocarbon liquids, including butanes, pentanes,hexanes, heptanes, octanes, aromatics, and their various isomers, bothsaturated (alkanes) and unsaturated (alkenes), or it can be any one ofthese hydrocarbon liquids. In a preferred embodiment, the solvent iscommercial grade normal-butane or a mixture of normal-butane andiso-butane.

The carbon dioxide is removed from the solvent using pressure andthermal changes and is produced at medium pressure (ranging from 200-500psig). In a preferred embodiment, this system uses externalrefrigeration to aid in the carbon dioxide absorption and low-pressuresteam for thermal regeneration of the solvent. Other heat mediums may beused to provide the heat to thermally regenerate the solvent.

Although described herein as the fourth system (after the inlethandling, syngas and water-gas shift, and water-gas compressionsystems), the carbon dioxide recovery system could be third (beforewater-gas compression) or fifth (after water-gas compression anddehydration) in the processing sequence. Thus, the order of the systemsis not critical. The operating pressures of the absorber in the carbondioxide recovery system may range from 0 psig to 2200 psig, dependingupon the application (for example, if the present invention were used torecover carbon dioxide from low-pressure sources, such as fromcombustion flue gas, coal-bed methane gas, or other carbon-dioxidebearing sources). The feed temperature for this system will typicallyrange from 90° F. to 120° F., depending upon the means used to cool thewater-gas feed stream to the carbon dioxide recovery system.

The carbon dioxide recovery system requires that the water-gas feedstream be dehydrated prior to or in conjunction with cooling the feedstream. In the preferred embodiment, the dehydration system uses asolution of ethylene glycol and water (“ethylene glycol solution”) todehydrate the water-gas feed stream. In alternate embodiments, both ofwhich are described in Section F (Dehydration System), either a methanolsolution, absorption (e.g., triethylene glycol), or an adsorption (e.g.,molecular sieve, activated alumina, or silica gel) dehydration system isutilized upstream of the carbon dioxide recovery system to remove thewater vapor from the water-gas.

The carbon dioxide system is comprised of two separate sections: the gasabsorption section and the carbon dioxide recovery solvent regenerationsection. The gas absorption section (shown in FIG. 5) uses the carbondioxide recovery solvent to absorb the carbon dioxide from thewater-gas. The carbon dioxide recovery solvent section (shown in FIGS. 6and 7) removes the carbon dioxide from the carbon dioxide recoverysolvent and recycles the carbon dioxide recovery solvent back to the gasabsorption section of the carbon dioxide recovery system.

Referring to FIG. 5, the gas absorption section of the carbon dioxidesystem is comprised of a gas/gas exchanger 28, a high-level gas chiller31, a low-level gas chiller 32, a glycol separator 34, a primaryabsorber 36, a presaturator separator 29, a presaturation chiller 33,two presaturator pumps 30 (so that the process can continue to run ifone pump breaks), and an ethylene glycol exchanger 35.

Referring to FIG. 6, the carbon dioxide recovery solvent regenerationsection of the carbon dioxide recovery system further comprises a flashabsorber 39, a flash gas compressor suction scrubber 38, a flash gascompressor 37, a high-level solvent chiller 41, and a low-level solventchiller 40. In a preferred embodiment, ethylene glycol is used inconnection with the dehydration system (described below), and thepresent invention comprises an ethylene glycol regeneration system. Thefollowing components of the ethylene glycol regeneration system (whichis designed to remove the absorbed water from the ethylene glycolsolution) are also shown on FIG. 6: an ethylene glycol heater 42 and anethylene glycol flash separator 43. Referring to FIG. 7, the carbondioxide recovery solvent regeneration section of the carbon dioxiderecovery system further comprises a solvent cross exchanger 44, asolvent stripper 46, a solvent stripper reflux condenser 49, a solventstripper reflux accumulator 50, two (in case one breaks down) solventstripper reflux pumps 47, a solvent stripper reboiler 48, two (in caseone breaks down) solvent charge pumps 45, and a refrigerant subcooler51. Referring to FIG. 8, the carbon dioxide recovery system furthercomprises a refrigeration system.

The ethylene glycol regeneration system (not shown in the figures exceptfor those parts shown on FIG. 6) is designed to remove the absorbedwater from the rich (i.e., water-bearing) ethylene glycol solution,converting it to lean (i.e., water-short) ethylene glycol solution;filter and clean the ethylene glycol solution; and pump the leanethylene glycol solution back to the carbon dioxide recovery system forreuse in dehydrating the water-gas. This system is a closed loop systemand is a common, industry-standard design in the gas processingindustry.

Referring to FIG. 5, the water-gas feed stream from the water-gascompression system enters the carbon dioxide recovery system at 90° F.and 1270 psia. As is typical for the output of a steam reformingreaction (shown in FIG. 3), this gas flow rate is 452 MM SCFD (based onan inlet feed rate of 100 MM SCFD to the water-gas shift system), thegas is saturated with water vapor and contains approximately 20% carbondioxide, 77% hydrogen and 3% other gases (by volume). The water-gas feedstream is cooled to −35° F. and 1255 psia in a series of threeexchangers: the gas/gas exchanger 28, the high-level gas chiller 31, andthe low-level gas chiller 32. The gas/gas exchanger 28 uses coldhydrogen-rich product residue gas (described below) from thepresaturator separator 29 to cool the gas stream from 90° F. to 13° F.The high-level gas chiller 31 uses external refrigeration to cool thegas stream from 13° F. to 0° F., and the low-level gas chiller 32 usesexternal refrigeration to cool the gas stream from 0° F. to −35° F. Thisgas cooling reduces the volume of carbon dioxide recovery solvent neededto remove the carbon dioxide from the syngas, thereby reducing overallcapital and operating costs.

In a preferred embodiment, three heat exchangers (one gas/gas exchangerand two refrigerant chillers) are used to minimize overall energyrequirements of the external refrigeration system. Other heat exchangerarrangements could also be employed without changing the unique aspectsof the carbon dioxide recovery system. This includes using refrigerationsystems that may lower the temperature of the water-gas feeding theabsorber below or above −35° F., depending on the optimization of thecapital and operating costs that are affected by the refrigerantsselected, carbon dioxide recovery solvent circulation rate, number oftheoretical trays of the absorber, and carbon dioxide recoveryefficiency. Lean ethylene glycol solution (a mixture of 80 weightpercent ethylene glycol and 20 weight percent liquid water) is injectedinto the inlet gas side of each exchanger pass to prevent hydrateformation; the resulting rich (a mixture of 70 weight percent ethyleneglycol and 30 weight percent liquid water) ethylene glycol solution isremoved in the glycol separator 34.

The cooled gas then enters the primary absorber 36, where it iscounter-currently contacted by carbon dioxide recovery solvent from thepresaturator separator 29 (described below). The carbon dioxide recoverysolvent absorbs the bulk of the carbon dioxide in the gas. Thisabsorption releases energy, which heats the water-gas in the primaryabsorber 36 as it travels up and out of the primary absorber. Thisreleased energy heats the overhead syngas stream exiting the top of theprimary absorber 36 from −35° F. to −18° F. and heats the rich carbondioxide recovery solvent (carbon dioxide-bearing carbon dioxide recoverysolvent) leaving the bottom of the primary absorber 36 from −35° F. to−8° F. The rich carbon dioxide recovery solvent then flows to the carbondioxide recovery solvent regeneration section of the carbon dioxiderecovery system for carbon dioxide recovery.

The overhead water-gas stream, which has become hydrogen-concentratedfrom the primary absorber 36, is mixed with approximately 3,200 gallonsper minute of lean carbon dioxide recovery solvent (regenerated carbondioxide recovery solvent with less than 5 weight percent carbondioxide). The absorption of the remaining carbon dioxide from theoverhead water-gas stream heats the mixture to −23° F. The mixture isthen cooled from −23° F. to −35° F. in the presaturation chiller 33 toremove the heat of absorption. The resulting mixture consists of a coldhydrogen-rich product gas and semi-lean carbon dioxide recovery solvent(carbon dioxide recovery solvent that carries carbon dioxide but has thecapacity to remove additional carbon dioxide from the water-gas). Thesemi-lean carbon dioxide recovery solvent is separated from this coldhydrogen-rich product gas in the presaturator separator 29 and pumped tothe primary absorber 36 by the presaturator pump 30. The coldhydrogen-rich product gas is heated to 75° F. against inlet gas in thegas/gas exchanger 28 and sent to the residue pipeline at 1220 psia orthe pipeline operating pressure. If required, a hydrogen compressor canbe utilized to inject the hydrogen-rich gas into the gaseous hydrocarbonstream.

The rich ethylene glycol solution that is separated from the chilledwater-gas feed stream in the glycol separator 34 absorbs carbon dioxideas well as water. To recover the bulk of the absorbed carbon dioxide,the rich ethylene glycol solution is heated to approximately 150° F. bycross exchange with hot lean ethylene glycol solution from the ethyleneglycol regeneration system in the ethylene glycol exchanger 35 (shown onFIG. 5) and with 50 psig steam (or other suitable heat medium) in theethylene glycol heater 42 (shown on FIG. 6), respectively. The hot richethylene glycol solution stream pressure is reduced from about 1220 psigto 435 psig by a level control valve (shown to the left of the ethyleneglycol flash separator 43 on FIG. 6), thereby reducing the temperatureto 81° F. and releasing the bulk of the absorbed carbon dioxide. Thetwo-phase stream flows to the ethylene glycol flash separator 43, fromwhich the released carbon dioxide flows to the carbon dioxidecompression system, while the rich ethylene glycol solution flows to theethylene glycol regeneration system for removal of the absorbed water.

The rich carbon dioxide recovery solvent leaving the primary absorber 36is reduced in pressure by a level control valve (shown to the bottomright of the primary absorber 36 on FIG. 5) from about 1230 psig to 435psig, releasing a large part of the hydrogen and methane gas that wasabsorbed in the primary absorber 36. This pressure reduction alsoreleases a portion of the absorbed carbon dioxide and vaporized aportion of the carbon dioxide recovery solvent. To minimize solvent andcarbon dioxide losses from this operation, this two-phase stream flowsto the flash absorber 39, where the released gas is counter-currentlycontacted by 356 gallons per minute of lean carbon dioxide recoverysolvent that has been cooled to −35° F. in the high-level solventchiller 41 (described below). The cold carbon dioxide recovery solventabsorbs the carbon dioxide and cools the released gas to −18° F. in theflash absorber 39.

Referring to FIG. 6, the overhead gas exiting the top of the flashabsorber 39 is hydrogen-rich and has minimal amounts of carbon dioxideand vaporized carbon dioxide recovery solvent present in the stream.This overhead gas stream flows to the flash gas compressor suctionscrubber 38 to remove any entrained liquids. Next, it is compressed to1225 psig by the flash gas compressor 37, which raises the temperatureof the gas is raised to 127° F. This gas mixes with the gas exiting thegas/gas exchanger 28, and the mixture then becomes the hydrogen productstream (see FIGS. 2 and 5). In an alternate embodiment, the flash gasmay be used as fuel within the present invention and therefore eliminatethe need to reduce the size of the flash gas compression system. Thishydrogen product stream mixes with the gaseous hydrocarbon bypass stream(previously described and shown on FIG. 2) to become the hydrogen-richgaseous hydrocarbon product. A lesser portion of this stream is divertedto the burner of the reformer 19 as fuel (FIGS. 2 and 3). The balance ofthe hydrogen-rich gaseous hydrocarbon product is delivered to end usersvia pipeline.

Referring to FIG. 7, the rich carbon dioxide recovery solvent from thebottom of the flash absorber 39 is heated to 175° F. by hot lean carbondioxide recovery solvent from the bottom of the solvent stripper 46 inthe solvent cross exchanger 44 before flowing to the solvent stripper46. The solvent stripper 46 is a full distillation column that removesthe absorbed carbon dioxide from the rich carbon dioxide recoverysolvent, thereby regenerating the solvent back to the carbon dioxidelean condition where it can then be fed back to the primary absorber 36,the presaturator chiller 33 and the flash absorber 39 for carbon dioxiderecovery. Heat for this distillation is supplied by the solvent stripperreboiler 48 at 271° F. and 400 psig, using 50 psig steam as the heatingmedium to vaporize the carbon dioxide and a portion of the carbondioxide recovery solvent from the total carbon dioxide recovery solventstream feeding the reboiler. Alternately, the heat could be suppliedfrom a heat medium system or a direct-fired heater.

Referring to FIG. 7, the overhead gas exiting the top of the solventstripper 46 then flows to the solvent stripper reflux condenser 49,where it is cooled to 15° F. by external refrigeration (shown in FIG.8), condensing most of the vaporized carbon dioxide recovery solventfrom the carbon dioxide, resulting in a carbon dioxide recovery stream(reflux) containing little carbon dioxide, and a gas stream (carbondioxide product gas) containing little carbon dioxide recovery solvent.The carbon dioxide recovery solvent is separated from the carbon dioxideproduct gas in the solvent stripper reflux accumulator 50 and returnedto the solvent stripper 46 by the solvent stripper reflux pump 47. Thecarbon dioxide product gas flows to the refrigerant subcooler 51, whereit is heated to 41° F. before flowing to the carbon dioxide compressionsystem with the flash gas from the ethylene glycol flash separator 43(described above). In an alternate embodiment, the stripper reflux maybe operated as a total condenser and the carbon dioxide pumped to therequired delivery pressure.

As stated above, in a preferred embodiment, the carbon dioxide recoverysolvent is commercial grade normal-butane or a mixture of normal-butaneand iso-butane. The physical characteristics of this hydrocarbon providegood absorption of the carbon dioxide and allow the solvent stripper 46to operate at between 200 and 500 psig (400 psig in a preferredembodiment), producing a carbon dioxide product gas at this pressure.This high pressure greatly reduces the carbon dioxide compression energyrequirements when compared to the present art, which typically does notprovide a carbon dioxide product gas that is higher than 10-15 psig.

The lean carbon dioxide recovery solvent leaves the solvent stripper 46at 271° F. and 415 psia and is cooled in the solvent cross exchanger 44to 34° F. by incoming rich carbon dioxide recovery solvent. The cool,lean carbon dioxide recovery solvent is then pumped to 1255 psia by thesolvent charge pump 45 at a flow rate of about 3,780 gallons per minute.The lean carbon dioxide recovery solvent is mixed with about 49 gallonsper minute of make-up solvent before flowing to the high-level solventchiller 41 and the low-level solvent chiller 40. The coolant for thesechillers is preferably an external propane refrigerant (shown on FIG.8), although other refrigeration systems may be utilized. The high-levelsolvent chiller 41 first reduces the solvent temperature from 40° F. to0° F., and the low-level solvent chiller 40 further reduces thetemperature from 0° F. to −35° F. In the preferred embodiment, twochillers are used to minimize overall energy requirements of theexternal refrigeration system; however, the present invention is notlimited to one, two or three-chiller designs or to any particular methodof refrigeration. After exiting the low-level solvent chiller 40, thischilled lean carbon dioxide recovery solvent is then split into threeseparate streams that flow to the primary absorber 36 and thepresaturator chiller 33 (both shown on FIG. 5) and the flash absorber 39(shown on FIG. 6).

The carbon dioxide recovery system is integrated with the externalrefrigeration system (shown in FIG. 8). The refrigeration system is aclosed-loop system that provides cooling for the various liquid and gasstreams in the carbon dioxide recovery system. Various industrialrefrigerants could be used for this service, including ethane, propane,butane, ammonia (both compression-cycle and absorption cycle designs),lithium bromide, carbon dioxide, fluorocarbons (Freons), fluorocarbonreplacements (e.g., R134a), mixed refrigerants or any other refrigerantavailable now or in the future that could provide refrigeration at thenecessary temperature levels.

In a preferred embodiment, the refrigeration system uses a standardpropane refrigerant to provide the necessary cooling services atsub-ambient temperatures. In general, liquid propane (refrigerant) fromthe refrigeration system is reduced in pressure, which reduces thetemperature of the refrigerant. This liquid is vaporized in heatexchangers (chillers), providing cooling for the various liquid and gasstreams in the carbon dioxide recovery system. The vaporized refrigerantis compressed to a pressure between 130 psig and 250 psig, depending onthe refrigerant condenser cooling media (air, water, or otherrefrigerants) and ambient conditions, then condensed back to a liquid byeither cooling water or by an air cooler. The cycle then repeats,providing continuous cooling to the various liquid and gas streams inthe carbon dioxide recovery system.

In a preferred embodiment, liquid refrigerant at 90° F. and 143 psig isstored in the refrigerant accumulator 58, which provides a liquid surgevolume for the closed loop system. The liquid refrigerant is reduced to75 psig through a control valve. The pressure reduction reduces thetemperature to 48° F. and vaporizes a lesser portion of the refrigerant.The resulting two-phase system is separated in the refrigerant firststage economizer 57, with the overhead refrigerant gas going to thethird stage refrigerant compressor 62.

The liquid refrigerant from the refrigerant first stage economizer 57 iscooled from 48° F. to 44° F. by cold carbon dioxide product gas from thesolvent stripper 46 in the refrigerant subcooler 51 before being splitinto four separate streams through four separate control valves to 19psig. The pressure reduction reduces the temperature to 5° F. andvaporizes a lesser portion of the refrigerant. Three of the resultingtwo-phase streams flow to high-level chillers described above: thehigh-level gas chiller 31 (shown on FIGS. 5 and 8); the high-levelsolvent chiller 41 (shown on FIGS. 6 and 8); and the solvent stripperreflux condenser 49 (shown on FIGS. 7 and 8). The refrigerant in each ofthese chillers is vaporized, providing cooling for the various liquidand gas streams in the carbon dioxide recovery system. These three gasrefrigerant streams and the fourth two-phase refrigerant stream flow tothe refrigerant second stage economizer 56 at 5° F. and 19 psig, wherethe gas refrigerant is separated from the liquid refrigerant. The gasrefrigerant from the refrigerant second stage economizer 56 flows to thesuction of the second stage refrigerant compressor 55.

The liquid refrigerant from the refrigerant second stage economizer 56is split into three separate streams through three separate controlvalves, which reduces the pressure of the liquid refrigerant to lessthan 1 psig. This pressure reduction reduces the temperature of theliquid refrigerant to −40° F. and vaporizes a lesser portion of therefrigerant. The three resulting two-phase streams flow to low-levelchillers described above: the low-level gas chiller 32 (shown on FIGS. 5and 8); the low-level solvent chiller 40 (shown on FIGS. 6 and 8); andthe presaturator chiller 33 (shown on FIGS. 5 and 8). The refrigerant ineach of these chillers is vaporized, providing cooling for the variousliquid and gas streams in the carbon dioxide recovery system. Thesethree gas refrigerant streams then flow to the refrigerant compressorfirst stage suction scrubber 53 at −40° F. and less than 1 psig, wherethe gas refrigerant is separated from any excess liquid refrigerant. Thegas refrigerant from the refrigerant compressor first stage suctionscrubber 53 flows to the suction of the first stage refrigerantcompressor 54.

F. Dehydration System

The dehydration system removes the water present in the water-gas stream(from the inlet of the carbon dioxide removal system) down to a levelthat is acceptable for the carbon dioxide recovery system and/or to meetfinal product delivery requirements. Although described herein as thefifth system, the dehydration step actually occurs in conjunction withthe absorption section of the carbon dioxide recovery system.

In a preferred embodiment, and as described above, the dehydrationprocess involves the use of ethylene glycol solution injected directlyinto the process stream. Alternate dehydration technologies may be used,however. One alternative is to use methanol in the same way as theethylene glycol by injecting the methanol directly into the processstream. Another alternate dehydration technology involves the use ofabsorption (triethylene glycol) to remove the water from the water-gasprior to entering the carbon dioxide recovery system. Yet anotheralternate dehydration technology involves the use of adsorption(molecular sieve, activated alumina or silica gel beds) to remove thewater from the water-gas prior to entering the carbon dioxide recoverysystem. With the preferred embodiment, however, the dehydration systemoccurs in conjunction with the carbon dioxide recovery system. Thepreferred embodiment is described below, after which the alternatedehydration technologies are discussed.

In the preferred embodiment, the dehydration system involves injectionof lean ethylene glycol solution into the water-gas entering the gas-gasexchanger 28, the high-level gas chiller 31, and the low-level gaschiller 32 of the absorption section of the carbon dioxide recoverysystem. The lean ethylene glycol solution absorbs the bulk of the watervapor from the water-gas, creating a two-phase stream of dry syngas andrich ethylene glycol solution. The two-phase stream is separated in theglycol separator 34. The dry syngas then flows to the primary absorber36.

The dehydration system is further comprised of an ethylene glycol heater42 (shown on FIG. 6), an ethylene glycol flash separator 43 (shown onFIG. 6), and an ethylene glycol regeneration system (not shown). Theseexchangers are described above in connection with the absorption sectionof the carbon dioxide recovery system.

As stated above, the ethylene glycol regeneration system is designed toremove the absorbed water from the rich (i.e., water-bearing) ethyleneglycol solution, converting it to lean (i.e., water-short) ethyleneglycol solution; filter and clean the ethylene glycol solution; and pumpthe lean ethylene glycol solution back to the carbon dioxide recoverysystem for reuse in dehydrating the water-gas. This system is a closedloop system and is a common, industry-standard design.

The ethylene glycol regeneration system provides the lean ethyleneglycol solution for injection into the gas/gas exchanger 28 (shown onFIG. 5), the high-level gas chiller 31 (shown on FIG. 5), and thelow-level gas chiller 32 (shown on FIG. 5). The lean ethylene glycolsolution is injected into the inlet of each pass of these three heatexchangers, mixed with the water-gas stream, and then cooled as thetwo-phase flow (i.e., liquid and gas) flows through the three heatexchangers in sequence: the gas/gas exchanger 28, the high-level gaschiller 31, and the low-level gas chiller 32. The two-phase flow fromthese three heat exchangers is cooled to −35° F. at 1255 psia. Throughthe process of mixing and cooling with the water-gas feed stream, thelean ethylene glycol solution absorbs the liquid water condensed fromthe water-gas stream and protects the carbon dioxide recovery systemfrom both free water and freezing problems. It also removes enough watervapor to meet downstream hydrogen product gas and carbon dioxide productgas specifications. The cold two-phase stream (gas and rich ethyleneglycol) is separated in the glycol separator 34 (shown on FIG. 5). Thewater-gas stream, which is now free of liquid water and the build of thewater vapor, then flows to the primary absorber 36 of the carbon dioxiderecovery system, and the rich ethylene glycol flows to the ethyleneglycol exchanger 35.

The ethylene glycol exchanger 35 is used to cool the warm (240° F.) leanethylene glycol solution with the cold (−35° F.) rich ethylene glycolsolution. This cools the lean ethylene glycol solution to −16° F. byheating the rich ethylene glycol, thus increasing the energy efficiencyof the dehydration system and reducing the refrigeration requirements.The cold lean ethylene glycol solution is then injected into the gas/gasexchanger 28, the high-level gas chiller 31, and the low-level gaschiller 32.

A minor amount of carbon dioxide is absorbed by the rich ethylene glycolsolution in the gas/gas exchanger 28 and the high-level and low-levelgas chillers 31, 32. To recover the absorbed carbon dioxide, the richethylene glycol solution is heated to 81° F. in the ethylene glycolheater 42) (shown on FIG. 6). The heat medium for this heater ispreferably low-pressure (50 psig) steam, although other heat mediums maybe used. The heated rich ethylene glycol solution pressure is reduced to435 psia via a level control valve upstream of the ethylene glycol flashseparator 43 (shown on FIG. 6). The bulk of the carbon dioxide absorbedin the rich ethylene glycol solution is released as vapor as a result ofthe increase in temperature and reduction of pressure of the richethylene glycol solution. The resulting two-phase rich ethylene glycolsolution then flows to the ethylene glycol flash separator 43, where thereleased carbon dioxide vapor is separated from the rich ethylene glycolsolution.

The carbon dioxide from the ethylene glycol flash separator 43 (shown onFIG. 6) flows to the carbon dioxide compression system, where it ismixed with the carbon dioxide product gas from the solvent stripperreflux accumulator 50 carbon dioxide recovery system prior to beingcompressed in the carbon dioxide compression system.

The rich ethylene glycol solution from the ethylene glycol flashseparator 43 (shown on FIG. 6) flows to the ethylene glycol regenerationsystem, where the excess water in the rich glycol stream is removed fromthe rich ethylene glycol solution to provide lean ethylene glycolsolution. The water from the rich ethylene glycol solution is recoveredand treated to supply make-up boiler feed water to the syngas andwater-gas shift system. The lean ethylene glycol solution is then pumpedup to the feed required injection pressure of 1270 psia.

As noted above, other alternate dehydration technologies may be utilizedin connection with the present invention. The first involves utilizingthe same basic dehydration design but changing the fluid from ethyleneglycol to methanol. The second involves the use of triethylene glycol.With this process, the feed gas would be contacted with the triethyleneglycol in a counter-current gas/liquid contactor. In this contactor, thefeed gas would flow into the bottom to the top of the contactor, whereit would be contacted by lean triethylene glycol solution flowing downthe contactor. The triethylene glycol would remove the water in the feedstream, and the water-rich triethylene glycol solution would flow out ofthe bottom of the contactor to a triethylene glycol regeneration system.The triethylene glycol regeneration system would strip the absorbedwater from the triethylene glycol solution, and the water vapor would bevented to atmosphere. The lean triethylene glycol would then berecirculated back to the triethylene glycol contactor.

The use of this type of dehydration would eliminate the ethylene glycolsolution injection in the absorption section of the carbon dioxiderecovery system. It would also eliminate the glycol separator 34, theethylene glycol heater 42, the ethylene glycol exchanger 35, and theethylene glycol flash separator 43, and their associated streams.

In an alternate dehydration technology, the process gas stream wouldenter multiple (two or more) absorption dehydrator beds. Thesedehydrator beds would be filled with dessicant, and they would operatein a cyclic manner. The selected desiccant would be molecular sieve,activated alumina, silica gel, or similar dessicant, and it would beporous with a high affinity for water.

In this process, the dessicant in the one or more dehydrator bed(s)would adsorb the water vapor from the water-gas, while the one or moredehydrator bed(s) that are saturated with water would be regenerated byheating and/or pressure reduction to remove the adsorbed water. Thecontrols would automatically adjust the feed gas rate to the properdehydrator bed to provide continual feed stream dehydration.

The water-saturated water-gas would flow downward through the dehydratorbed in the adsorption cycle. The water would be adsorbed within thepores of the dessicant, thereby drying the water-gas stream to themoisture content suitable for gas processing in the carbon dioxiderecovery system. Once the desiccant in the dehydrator bed in theadsorption cycle is saturated with water, it would be taken off-line andthe water-gas stream would be diverted to other dehydrator beds. Thesaturated dehydrator bed would be regenerated first by heating it withina preheated portion of the dehydrated water-gas leaving the dehydrationsystem (called “regeneration gas”) at either operating pressure or lowpressure to drive the adsorbed water off the dessicant into theregeneration gas, making it suitable for reuse in the next cycle, andthen cooling it with unheated regeneration gas.

If this dehydration technology were used in connection with the presentinvention, the regeneration gas exiting the dehydration system could berecycled back into the inlet of the carbon dioxide recovery process.

G. Carbon Dioxide Compression System

FIG. 9 is a flow diagram of the carbon dioxide compression system of thepresent invention. The carbon dioxide compression system is designed tocompress the various recovered carbon dioxide product stream from thecarbon dioxide recovery system to the pressure required by the carbondioxide storage/sequestration or EOR system. The carbon dioxidecompression system inlet pressures may range from 0 psig to 500 psig,depending upon the configuration of the carbon dioxide recovery system,The type of compressor used in this system e., centrifugal,reciprocating, etc.) will depend upon the flows and pressures requiredfor the carbon dioxide sequestration or EOR system, equipment sizing,and manufacturer selection/pricing.

Referring to FIG. 9, the carbon dioxide compression system is comprisedof a carbon dioxide compressor suction scrubber 59, a carbon dioxidecompressor 60, and a carbon dioxide compressor discharge cooler 61.

In a preferred embodiment, the carbon dioxide product stream from thecarbon dioxide recovery system and the carbon dioxide flash gas from thedehydration system are combined and flow to the carbon dioxidecompressor suction scrubber 59. The suction scrubber removes anyentrained liquids (for example, ethylene glycol or carbon dioxiderecovery solvent), and these liquids flow to the high-pressure drainsystem (not shown). The carbon dioxide gas stream flows to the carbondioxide compressor 60, where the stream is compressed to 2220 psig orother required delivery pressure established for carbon dioxidetransportation, storage, EOR, or other sales. The compressed carbondioxide stream is then cooled to 120° F. in the carbon dioxidecompressor discharge cooler 61 prior to flowing into the carbon dioxidesequestration or EOR system. Alternatives to utilizing a compressor toincrease the carbon dioxide pressure to the required delivery conditionsinclude employing a pump.

The specific configurations and operating temperatures, pressures andflow rates discussed above are provided for illustrative purposes onlybut are not intended to limit the scope of the present invention.Although the preferred embodiment of the present invention has beenshown and described, it will be apparent to those skilled in the artthat many changes and modifications may be made without departing fromthe invention in its broader aspects. The appended claims are thereforeintended to cover all such changes and modifications as fall within thetrue spirit and scope of the invention.

REFERENCES

1. “Causes of Global Warming.” EcoBridge.http://www.ecobridge.org/content/g_cse.htm.

2. “Climate Change 2007.” The Fourth Assessment Report of the UnitedNations Intergovernmental Panel on Climate Change.http://en.wikipedia.org/wiki/IPCC_Fourth_Assessment_Report.

3. “Alternative & Advanced Fuels.” U.S. Department of Energy (Sep. 18,2007). http://www.eere.energy.gov/afdc/fuels/natural_gas_blends.html.

4. “Program Overview: Hydrogen Enriched Compressed Natural Gas.”Westport Innovations Inc. (December 2005).http://www.westport.com/pdf/WPT-HCNG_MED.pdf.

5. “Glossary of Terms.” Virginia Department of Mines Minerals and Energy(2006). http://www.dmme.virginia.gov/DE/glossaryterms.shtml.

6. Raabe, Steve. “Research Going Underground.” The Denver Post (Feb. 18,2007) (citing Colorado Geological Survey, partner in Department ofEnergy's Southwest Regional Partnership on CO₂ Sequestration).

1. A method of generating hydrogen-enriched fuel gas and carbon dioxidecomprising: (a) converting hydrocarbon molecules from a gaseoushydrocarbon feed stream into hydrogen and carbon dioxide; (b) separatingthe hydrogen and carbon dioxide; (c) blending the hydrogen back into thegaseous hydrocarbon feed stream to generate a hydrogen-enriched fuelgas; and (d) utilizing the carbon dioxide for storage or sequestration.2. The method of claim 1, wherein as a result of the separation step,each standard cubic foot of the gaseous hydrocarbon feed. streamproduces between two and four standard cubic feet of a hydrogen productstream and between 0.7 and 0.9 standard cubic feet of a carbon dioxideproduct stream.
 3. The method of claim 1, wherein the hydrogen-enrichedfuel gas has a hydrogen concentration ranging from five to 30 molepercent.
 4. The method of claim 1, wherein the hydrogen-enriched fuelgas produces less carbon dioxide per energy unit output when combustedthan non-hydrogen-enriched natural gas.
 5. The method of claim 1,wherein approximately one to eleven percent of the gaseous hydrocarbonfeed stream is processed.
 6. The method of claim 1, wherein the gaseoushydrocarbon feed stream has a total volume; wherein the total volume ofthe gaseous hydrocarbon feed stream ranges from 100 million standardcubic feet per day to 4500 million standard cubic feet per day; andwherein between 10 million standard cubic feet per day and 500 millionstandard cubic feet per day of the gaseous hydrocarbon feed stream isprocessed.
 7. The method of claim 1, wherein there is an existingnatural gas pipeline transportation and distribution system, and whereinthe hydrogen-enriched fuel gas is transported and distributed using theexisting natural gas pipeline system.
 8. The method of claim 1, whereinthe sequestration is enhanced oil recovery.
 9. The method of claim 1,wherein the carbon dioxide is separated from the hydrogen using a carbondioxide recovery solvent, and wherein the carbon dioxide recoverysolvent is one or more hydrocarbon liquid(s) selected from the groupconsisting of butanes, pentanes, hexanes, heptanes, octanes, aromatics,and isomers of butanes, pentanes, hexanes, heptanes, octanes andaromatics.
 10. The method of claim 1, wherein a carbon dioxide recoverysolvent is used to separate the carbon dioxide from the hydrogen; andwherein the carbon dioxide recovery solvent allows the carbon dioxide tobe separated from the hydrogen at a pressure of between 200 and 500psig.
 11. The method of claim 10, wherein total carbon dioxidecompression requirements for the storage or sequestration are reduced by50 to 75 percent as compared to chemical or physical solvent-basedcarbon dioxide recovery processes that do not utilize the carbon dioxiderecovery solvent of the present invention.
 12. The method of claim 1,wherein the conversion, separation and blending steps occur on a naturalgas transportation and distribution pipeline and not at a point ofcombustion.
 13. The method of claim 1, wherein the conversion,separation and blending steps are all performed prior to combustion ofthe hydrogen-enriched fuel gas.
 14. The method of claim 1, wherein thereis an existing natural gas pipeline transportation and distributionsystem, wherein there are a number of existing compressed natural gasfueling facilities; and wherein the conversion, separation and blendingsteps do not require any changes to the existing natural gas pipelinetransportation and distribution system other than providing mobilepoints of consumption with an ability to consume the hydrogen-enrichedfuel gas and increasing the number of compressed natural gas fuelingfacilities to supply the mobile points of consumption with thehydrogen-enriched fuel gas.
 15. The method of claim 1, furthercomprising: (e) utilizing a portion of the separated hydrogen as aseparate fuel product.
 16. The method of claim 1, wherein the gaseoushydrocarbon feed stream is pipeline quality natural gas.
 17. A systemfor generating hydrogen-enriched fuel gas and carbon dioxide comprising:(a) means for converting hydrocarbon molecules from a gaseoushydrocarbon feed stream into hydrogen and carbon dioxide; (b) means forseparating the hydrogen and carbon dioxide; (c) means for blending thehydrogen back into the gaseous hydrocarbon feed stream to generate ahydrogen-enriched fuel gas; and (d) means for utilizing the carbondioxide for storage or sequestration.
 18. A method of generatinghydrogen and carbon dioxide comprising: (a) converting hydrocarbonmolecules from a gaseous hydrocarbon feed stream into hydrogen andcarbon dioxide; (b) separating the hydrogen and carbon dioxide; (c)utilizing the hydrogen as a separate product; and (d) utilizing thecarbon dioxide for storage or sequestration; wherein the carbon dioxideis separated from the hydrogen using a carbon dioxide recovery solvent;and wherein the carbon dioxide recovery solvent is one or morehydrocarbon liquid(s) selected from the group consisting of butanes,pentanes, hexanes, heptanes, octanes, aromatics, and isomers of butanes,pentanes, hexanes, heptanes, octanes and aromatics.
 19. The method ofclaim 18, wherein a carbon dioxide recovery solvent is used to separatethe carbon dioxide from the hydrogen; and wherein the carbon dioxiderecovery solvent allows the carbon dioxide to be separated from thehydrogen at a pressure of between 200 and 500 psig.